Batch process and system for the production of olefins

ABSTRACT

Disclosed herein is a process for producing an alpha olefin comprising obtaining a feed stream comprising an internal olefin having a first carbon number and an alpha olefin having a first carbon number, isomerizing the feed stream to increase the quantity of the alpha olefin, fractionating, subjecting the overhead material from fractionation to catalytic metathesis to produce a mixed olefin effluent comprising an internal olefin having a second carbon number and other hydrocarbons, fractionating, preparing the first isomerization reactor and fractionator to receive the olefin having a second carbon number, isomerizing the internal olefin intermediate in the prepared first isomerization reactor, and fractionating the second isomerization effluent in the prepared first fractionator to separate the alpha olefin having the second carbon number from the internal olefin having the second carbon number. A corresponding system also is disclosed, along with a heat pump that can be incorporated into the process.

CROSS-REFERENCE TO RELATED APPLICATIONS

This application, pursuant to 35 U.S.C. §120, claims benefit to U.S.application Ser. No. 12/936,284, now U.S. Pat. No. 8,742,186, issuedJun. 3, 2014, which is a national stage application claiming benefit toPCT/US2009/002076, filed Apr. 2, 2009, which pursuant to 35 U.S.C.119(e) claims priority to U.S. Provisional Application Ser. No.61/072,993, filed Apr. 4, 2008. Each of these applications isincorporated herein by reference in its entirety.

BACKGROUND

The disclosed embodiments generally relate to processes and systems forproducing alpha olefins and more particularly to a batch process for theproduction of alpha olefins.

A conventional process for production of comonomer grade hexene-1 fromC₄ raffinate feed streams is a continuous process that has three stages.First butene-1 is separated from the feed stream in a C₄ fractionator.The butene-2 in the fractionator bottoms stream is isomerized tobutene-1 and recycled to the fractionator. Second, the butene-1 is sentto an autometathesis reactor to form ethylene and hexene-3. The reactoreffluent is sent to a depentanizer to separate hexenes. The products arelights that go overhead, the hexene-3 is a liquid bottoms product, andthe C₄/C₅ products are recycled. Third, the hexene-3 feed is isomerizedand the hexene-1 product is separated in a C₆ fractionator.

U.S. Pat. No. 6,727,396 (Gartside, April 2004) describes a continuousprocess for production of hexene-1, combining the isomerization andmetathesis steps. Typical metathesis reactions are described in U.S.Pat. No. 3,595,920 (Ellis et al, July 1971). U.S. Pat. No. 4,709,115(Jung et al, November, 1987) discusses improving the selectivity andconversion of butene-1 and butene-2 to hexene-3 by using catalyticdistillation. The removal of the lighter components pushes the reactionequilibrium toward the heavy products. U.S. Pat. No. 5,057,638 (Sweeney,October 1991) discusses a method for production of hexene-1 frombutene-1 in which the butene-1 is metathesized to hexene-3.Subsequently, a hydration/dehydration procedure is applied to produce amixture of n-hexenes containing hexene-1.

Various other processes are known for the processing of C₄ olefins. U.S.Pat. No. 6,875,901 (Gartside et al, April 2005) describes olefinisomerization technology used for production of terminal olefins. Theprocess is applied to the production of butene-1 from butene-2. U.S.Pat. No. 6,777,582 (Gartside et al, August 2004), describes butene-1autometathesis technology, including differences from the conventionalmetathesis reaction of butene-2 and ethylene to produce propylene.

Closed-loop heat pumps are used in various processes. U.S. Pat. No.6,589,395 describes a process in which a closed-loop heat pump isincluded on a general distillation tower. This document describes theuse of a heat source and heat sink that can be substituted for the heatpump should the compressor fail. U.S. Pat. Nos. 5,386,075 (Keil et al,January 1995) and No. 4,615,769 (Horigome et al, October 1986) discussthe use of an open-loop heat pump in an ethylbenzene/styrenedistillation.

It would be useful to develop a process for producing alpha olefins thathas improved efficiency when operated on a small scale.

SUMMARY

One embodiment is a process for producing an alpha olefin comprisingobtaining a feed stream comprising an internal olefin having a firstcarbon number and an alpha olefin having a first carbon number,isomerizing the feed stream in a first isomerization reactor to increasethe quantity of the alpha olefin having the first carbon number, fanninga first isomerization effluent, fractionating the first isomerizationeffluent in a first fractionator to obtain a bottoms stream comprisingthe internal olefin having the first carbon number and an overheadstream comprising the alpha olefin having the first carbon number,subjecting the overhead stream to catalytic metathesis in a metathesisreactor under conditions and in the presence of a first metathesiscatalyst to produce a mixed olefin effluent comprising an internalolefin having a second carbon number and other hydrocarbons,fractionating the mixed olefin effluent in a second fractionator toremove at least a portion of the other hydrocarbons and obtain aninternal olefin intermediate, preparing the first isomerization reactorto receive the internal olefin intermediate, isomerizing the internalolefin intermediate in the prepared first isomerization reactor to forma second isomerization effluent comprising an increased quantity ofalpha olefins having the second carbon number, preparing the firstfractionator to receive the second isomerization effluent, andfractionating the second isomerization effluent in the prepared firstfractionator to separate the alpha olefin having the second carbonnumber from the internal olefin having the second carbon number. In someembodiments, a portion of the butene-1 is removed from the firstfractionator as butene-1 product.

Another embodiment is a process for producing hexene-1 comprisingobtaining a C₄ feed containing butene-1 and butene-2, isomerizingbutene-2 to butene-1 in a first isomerization reactor, forming a firstisomerization reactor effluent, fractionating the first isomerizationreactor effluent in a first fractionator to form an overhead streamcomprising butene-1 and a bottoms stream comprising butene-2, subjectingat least a portion of the overhead product to catalytic metathesis in afirst metathesis reactor under conditions and in the presence of a firstmetathesis catalyst to produce a mixed olefin effluent comprisingethylene and hexene-3, fractionating the mixed olefin effluent in asecond fractionator to form a hexene stream comprising hexene-3 and anoverhead product stream comprising ethylene, preparing the firstisomerization reactor to receive the hexene stream, isomerizing thehexene stream to form a second isomerization effluent comprisinghexene-1 and hexene-2 and the remaining hexene-3, preparing the firstfractionator to receive the second isomerization effluent, andfractionating the second isomerization effluent in the preparedfractionator to obtain a hexene-1 stream.

Yet another embodiment is a system for producing an alpha olefin,comprising a first isomerization reactor configured to isomerize a firstbatch of an olefin having a first carbon number to form a firstisomerization reactor effluent and subsequently process a second batchof an olefin having a second carbon number to form a secondisomerization reactor effluent, a metathesis reactor positioneddownstream from the first isomerization reactor, the metathesis reactorbeing configured to disproportionate the first isomerization reactoreffluent to form a metathesis reaction product, a first fractionatorpositioned downstream from the isomerization reactor and beingconfigured to separately fractionate the first and second isomerizationreactor effluents, a second fractionator positioned downstream from themetathesis reactor to remove light hydrocarbons from the metathesisreaction product, a storage tank disposed downstream from the first orsecond fractionator, and a storage tank outlet line connecting thestorage tank to an inlet of the first isomerization reactor and/or tothe inlet of the metathesis reactor.

BRIEF DESCRIPTION OF THE DRAWINGS

FIG. 1 is a schematic drawing showing three sections of the systemdescribed herein.

FIG. 2 is a process flow diagram showing a first embodiment.

FIG. 3 is a process flow diagram showing a second embodiment.

FIG. 4 is a process flow diagram showing the first fractionator with aclosed loop heat pump system capable of operating with both the firstisomerization effluent and the second isomerization effluent in theprocess of FIG. 3.

FIG. 5 is a graph showing the temperature profile of a C₄ fractionatoraccording to the embodiment of Example 3.

FIG. 6 is a graph showing the temperature profile of a depentanizeraccording to the embodiment of Example 3.

FIG. 7 is a graph showing the temperature profile of a first C₆fractionator according to the embodiment of Example 3.

FIG. 8 is a graph showing the temperature profile of a second C₆fractionator according to the embodiment of Example 3.

DETAILED DESCRIPTION

The embodiments described herein employ a process operated in a campaignor sequential processing mode with a single isomerization reactor, asingle superfractionator following the isomerization, and one or moremetathesis reactors with subsequent fractionation to obtain intermediateolefins streams, and to obtain a desired olefin product or products. Theseparation of closely boiling double bond isomers of any single carbonnumber requires significant energy and equipment. By using a singlesuperfractionator (or set of 2 superfractionators) to separate isomershaving a first carbon number in a first separation process and to thensubsequently use the same superfractionator (or set ofsuperfractionators) to separate isomers having a second carbon number ina second separation process, certain efficiencies can be realized.Similarly, by using a single isomerization reactor to isomerize acompound having a first carbon number in a first isomerization processand a compound having a second carbon number in a second isomerizationprocess, processing advantages will be achieved. The process can be usedwith feed streams having carbon chains with a variety of carbon numbersto produce product streams having desired carbon numbers. The process isparticularly useful for producing alpha olefins.

FIG. 1 illustrates a system that includes an isomerization andfractionation section 2, a metathesis and fractionation section 3 and astorage section 4. While the descriptions of FIGS. 1-8 refers to C4 andC6 hydrocarbons, hydrocarbons with other carbon numbers also can beprocessed in the systems that are described. Afractionator/isomerization reactor combination, designated as 2 andtermed the “superfractionator system”, first operates in C₄ service.Mixed C₄'s are introduced at 1 and are isomerized and then fractionatedat 2 to form a butene-1 isomerization product. The butene-1isomerization product is fed continuously at 5 to the metathesis andfractionation section 3 in which metathesis takes place. The metathesisreactor effluent is fractionated to form light products includingethylene and a hexene-3 product which is fed at 6 to a storage tank at4. When sufficient hexene-3 has accumulated in the storage tank, theisomerization and fractionation section 2 is prepared for alternateservice. The hexene-3 from the tank is then sent at 7 to theisomerization and fractionation section 2 system now in C₆ service,where the hexene-3 is isomerized and fractionated to form the hexene-1product, which is removed at 8.

In another configuration, the mixed C₄'s are processed in theisomerization and fractionation system 2 to form butene-1. The butene-1stream is sent at 9 to the storage section 4. When sufficient butene-1has accumulated, the isomerization and fractionation system 2 isprepared for alternate service. A portion of the butene-1 optionally canbe removed as a product and the remaining portion is fed at 10 to themetathesis and fractionation section 3. The metathesis reactor effluentis fractionated to produce light products including ethylene and ahexene-3 stream. The hexene-3 stream is then sent at 11 to theisomerization and fractionation section 2 where the hexene-3 isisomerized and the mixed hexene stream fractionated to form hexene-1product, which is removed at 8.

In all embodiments, all or part of the internal olefin stream from thebottom of the superfractionator separation may be recycled to theisomerization reactor to produce more butene-1 or hexene-1.

In a larger scale conventional, continuous autometathesis process,separate C₄ and C₆ systems are employed, allowing heat integrationbetween the systems to reduce utilities. For the campaign operationsystems described herein, an alternate means of reducing utility costsis used to achieve savings. More specifically, in certain embodiments, aheat pump is included in a campaign system designed to produce olefinssuch as polymer-grade hexene-1. The heat pump provides a heat-integratedfractionator, whereby the tower's condenser and reboiler share a commonheat transfer fluid. An open-loop heat pump uses the tower overheadstream as the heat transfer fluid. A closed-loop heat pump uses analternate fluid. The alternate fluid is chosen based upon the specificthermodynamic properties to allow for condensing and reboiling duties tobe achieved within reasonable pressures such that compression duties areminimized. For systems operating in campaign mode, the choice ofalternate fluid is especially advantageous since it must operate toachieve condensing and reboiling duties in the fractionation of twodifferent carbon numbers.

Referring to FIG. 2, a process flow diagram for a campaign process forsequentially producing butene-1 and hexene-1 is shown. The overallprocess is designated as 12. One portion of the equipment is used in C₄service only, a second portion of equipment is used in C₆ service only,and a third set of equipment is shared between both services.

The butene separation system consists of two towers operated withdifferent pressures to allow for energy interchange between them toreduce overall utilities. Tower 14 is considered tower 1 and tower 24 isconsidered tower 2. Tower 1 operates at a higher pressure than tower 2.This allows the temperature of tower 1 overhead condenser 17 to be at ahigher temperature than tower 2 reboiler. Since heat is removed in thecondenser 17 and supplied to a reboiler 86, these can be exchangedwithout separate external heat being required. The key to this system isto balance the duties to allow them to be matched. This matching isconventionally done by bypassing a side draw from one tower to theother. Optimally however this is done by splitting the main feed to thetower with the proportion to each tower adjusted such that the exchangerduties can be matched. The main feed from the isomerization section 47is split into line 19 to tower 1 and 27 to tower 2.

A C₄ raffinate in feed line 13, which contains butene-1 and butene-2,and usually also contains other C₄ hydrocarbons, enters the lower end offractionator 24 in which butene-1 and butene-2 are separated. Thebottoms line 15 from fractionator 24, which primarily contains butene-2,combines with line 25 (line 32 is not used in the C₄ processing phase)to form line 34, which enters the isomerization reactor loop, describedbelow. The effluent of this loop in line 47 is split into line 19, whichenters the middle of a fractionator 14, and line 27, which enters themiddle of fractionator 24. In fractionator 14 an overhead product ofbutene-1 is taken in overhead line 16. The material in line 16 iscondensed in a condenser 17, separated into a reflux line 29 for thefractionator 14 and a feed line 31 for the fractionator 13, in whichfurther separation of butene-1 and butene-2 takes place.

Fractionator bottoms line 25 is removed from the bottom of thefractionator 14 and combined with line 15 as indicated above. Afractionator reboiler line 20 removes material at the bottom of thefractionator 14. A purge line 18 is taken off the fractionator reboilerline 20 to prevent buildup of any heavy hydrocarbons in the towerbottoms. The remainder of the fractionator bottoms in line 21 arereboiled in reboiler 23 and returned to fractionator 14 where theyundergo separation.

Feed line 31 enters the fractionator 24 above the point of entry of feedline 27. Butene-1 is removed from the top of fractionator 24 in line 33and butene-2 is removed from the bottom of fractionator 24 in line 37.The top line 33 is condensed in a condenser 39 and is divided into areflux line 35 and line 48.

In the isomerization loop, the material in isomerization line 34 isvaporized in heat exchanger 36 and heated in heat exchanger 38 and thenfed to a furnace 40. Vaporized line 42 from the furnace 40 is fed to anisomerization reactor 44 in which some of the butene-2 is isomerized toform butene-1. The C₄ effluent from the reactor 44 leaves atbutene-1/butene-2 equilibrium. The reactor temperature defines theequilibrium and thus controls the composition. The reactor effluent line47 is cooled in heat exchanger 38 and sent to the fractionator 14. It isapparent to one skilled in the art that if the C₄ feed line 13 containsbutene-1 above the equilibrium level set by the isomerization reactorconditions, the feed line would be first sent to the tower 14 and thebutene-1 content recovered overhead with the butene-2 being fed to theisomerization reactor 44. Alternately if the composition of C4 feed line13 had little or no 1 butene, it could first be fed directly to theisomerization system.

Downstream from fractionator 24, the contents of line 48 are either sentto tank 41, or to another storage tank, where they are collected untilready for metathesis, or they are sent directly to the metathesissection for further processing, in which case tank 41 is not required.In FIG. 2, line 48 is shown as providing for flow both into and out ofthe storage tank 41. When metathesis is to take place, line 48 iscombined with a recycle line 56 containing C₄/C₅ to form anautometathesis feed line 58, which is fed to an autometathesis reactor52. Before metathesis, line 58 is vaporized in a heat exchanger 60,further heated in a heat exchanger 62, and then heated to reactiontemperature in an autometathesis furnace 64. The contents of line 58 arethen fed to the autometathesis reactor 52. Autometathesis is anequilibrium reaction in which hexene-3 is produced. Small amounts ofside products propylene, pentene-2, 2-methyl-pentene-2, and some C7salso are produced. In addition, a small amount of reverse isomerizationof butene-1 to butene-2 occurs. Of these side products only2-methyl-pentene-2, formed from the reaction of butene-1 withisobutylene, unfavorably affects the hexene-1 product purity because itboils lower than hexene-1 and is thus carried out with the overheadproduct of the final C₆ separation. Thus, the isobutylene content in theC₄ raffinate feed is required to be minimized to a level consistent withthe desired hexene-1 specification.

The autometathesis effluent in line 65 is a mixture of C₂s through C₇s.The contents of line 65 are cooled in heat exchanger 62 to form adepentenizer feed line, which is sent to a fractionator 70 (operatinghere as a depentenizer). A C₂/C₃ overhead line 71 is removed from thefractionator 70. The overhead line 71 is condensed in a condenser 74 anddivided into a reflux line 73 and a feed line 75 for a fractionator 77,which is operated here as a depropylenizer. The bottoms line 83 offractionator 70 is sent to a C₆ storage tank 84 where it is held untilthe fractionation and isomerization equipment is ready for C₆processing, or, if the equipment is ready, the bottoms line 83 proceedsdirectly to the fractionation and isomerization section as line 32 andtank 84 is not required. It is noted that one of the fractionators 24and 14 can be run as the depentenizer 70 if the C₄ is stored at anappropriate time to allow for reconfiguration of the fractionator.Alternately fractionator 70 can be configured through piping to addadditional fractionation stages to fractionators 14 and 24 thusproviding additional flexibility for processing and capital and energysavings by allowing fractionators 14 and 24 to be slightly smaller.

The top line 79 from the fractionator 77 is divided into a C2/C3 line80, which is sent to a steam cracker separation system, for example, anda reflux line 81. Line 82 is a side draw product line that is optionallyinstalled to allow for recovery of a higher purity 1 butene stream fromthe unreacted 1 butene. A side draw 85 is removed and passed through acooler to partially condense the vapor stream in the tower at that pointwhen it is reintroduced. By cooling at a temperature consistent withcooling water at this point, the refrigeration commonly used in theoverhead condenser can be reduced. The bottoms line 56 from thefractionator 77 contains C₄ and C₅ compounds and is combined with line48 to form the metathesis feed stream.

Either before metathesis (but after C₄ isomerization and fractionation)or after production of a sufficient amount of hexene-3, theisomerization reactor 44 and fractionators 24 and 14 are prepared for C₆service. The autometathesis reactor is not used in the second campaignif it was used in the first step to produce 3-hexene. The hexene-3 isfed from storage tank 84 in line 32. Line 32 becomes isomerization feedline 34. Line 34 is vaporized in heat exchanger 36, heated in heatexchanger 38, further heated in furnace 40, and fed as line 42 first tothe isomerization reactor 44. The reactor effluent in line 47 is sent tothe fractionators 24 and 14. The isomerization reactor 44, fractionator24 and fractionator 14 are now operating in C₆ service.

In C₆ service, the fractionator 14 bottoms line 20 of C₇ ⁺ is partiallypurged from the system in line 18, and the remainder is reboiled inreboiler 23 and returned to the fractionator as line 21. A side draw 25of hexene-2 and hexene-3 is taken from a lower stage of the fractionator14 (with the top defined as stage 1), combined with bottoms line 15 fromfractionator 24, partially purged in line 28, and partially recycled tothe isomerization reactor 44 in line 30. Line 30 combines with freshhexene-3 feed 32 from the C₆ storage tank 84 to form isomerizationreactor feed 34 (now operating in C₆ service). The overhead line 16 fromfractionator 14 is divided into a reflux line 29 and a feed line 31 forthe fractionator 24.

In the fractionator 24, hexene-1 is taken as overhead product in line33. Line 33 is divided into reflux line 35 and line 48. Line 48 containsthe hexene-1 product and is sent to tank 41. It is noted that themetathesis reactor is not involved in the processing of the C₆ line.

As indicated above, the shared equipment from the batch processflowsheet is designed for operation in both C₄ and C₆ service. Dependingon the type of equipment, this can be handled in different ways. Heatexchangers, for example, may have varying temperature approaches, butthe heat exchanger surface area may be adjusted by using multipleshells. Reactor capacity can be addressed by using multiple reactors.Because the fractionator towers cannot be handled in the same way as theheat exchangers or reactors, their design is chosen to remain fixedbetween services.

In order to use the same tower or towers as both the C₄ and C₆fractionator, and if the second tower is used as both the depentanizerand as part of a C₆ fractionator, the tower sizing must be identical forthe chosen flow rates. Because campaign operation allows for independentvariation of the flow rates between C₄ and C₆ service, operating timecan be used as a variable to adjust the flow rates such that a netyearly production of, e.g., 5 KTA hexene-1 is achieved. Using thisapproach, in one embodiment the C₄ process is operated for 2,000 hoursand then the C₆ process is operated for 5,333 hours.

Overall, the shared use of the fractionator and isomerization systemcomponents in the batch process eliminates 35 of the 64 pieces ofequipment from the continuous hexene-1 process. The continuous processhas 2 complete superfractionator/isomerization reactor systems comparedto just one for the campaign operation. The estimated reduction in totalinstalled capital cost is about 35-45%. This makes campaign operationespecially suited for smaller capacity installations. An example of aprocess using the configuration of FIG. 1 is provided below as Example1.

Referring to FIG. 3, a process flow diagram for another campaign processfor sequentially producing butene-1 and hexene-1 is shown. The overallprocess is designated as 110. One portion of the equipment is used in C₄service only, a second portion of equipment is used in C₆ service only,and a third set of equipment is shared between both services.

A C₄ raffinate in feed line 112, which contains butene-1 and butene-2,and usually also contains other C₄ hydrocarbons, combines with thecontents of line 134 and enters the isomerization reactor loop. Theeffluent of this loop in line 147 enters the middle of a fractionator114. In fractionator 114 an overhead product of butene-1 is taken inoverhead line 116. The contents of line 116 are condensed in a condenser117.

Fractionator bottoms in line 122 are removed from the bottom of thefractionator 114, partially purged in a purge line 128, and theremaining material in line 130 is combined with the contents of C₄ feedline 112 (line 132 is not in use during the C₄ phase) to formisomerization feed line 134. Purge line 128 is provided to remove anyn-butanes in the C₄ feed 112 that would accumulate in the system. Afractionator reboiler line 120 removes material at the bottom of thefractionator 114. A purge line 118 is taken off the fractionatorreboiler line 120 to prevent buildup of any heavy hydrocarbons in thetower bottoms. The remainder of the fractionator bottoms in line 121 isreboiled in reboiler 123 and returned to fractionator 114 where itundergoes separation.

The contents of isomerization line 134 are vaporized in heat exchanger136 and heated in heat exchanger 138 and then fed to a furnace 140.Vaporized material in line 142 from the furnace 140 is fed to anisomerization reactor 144, which in one embodiment is an equilibriumreactor operating at 343 C and 2978 kPa. The C₄ effluent from thereactor in this embodiment leaves at butene-1/butene-2 equilibrium withan approximate butene-1 concentration of 21%. The reactor effluent inline 147 is cooled in heat exchanger 138 and sent to the fractionator114. The ratio of recycle to fresh feed in the fractionator is typically2.4 to 1. It is apparent to those skilled in the art that if the C₄ feedin line 112 contains butene-1 above the equilibrium level set by theisomerization reactor conditions, the feed stream would be first sent tothe tower 114 and the butene-1 content recovered overhead with thebutene-2 being fed to the isomerization reactor 144.

The butene-1 product from the fractionator 114 in overhead line 116 isdivided into a reflux stream in line 119 and intermediate product inline 148. The material in line 148 is sent into line 150 (line 154 isused for C₆ processing). The butene-1 in line 150 is combined with arecycle line 156 containing C₄/C₅ to form an autometathesis feed line158, which is fed to an autometathesis reactor 152. The material in line158 is vaporized in a heat exchanger 160, further heated in a heatexchanger 162, and then heated to reaction temperature in anautometathesis furnace 164. Vaporized material in line 159 is then fedto the autometathesis reactor 152. Autometathesis is an equilibriumreaction

The autometathesis effluent in line 166 is a mixture of C₂s through C₇s.The material in line 166 is cooled in heat exchanger 162 to form a(depentenizer) feed line. The contents of line 170 are sent to afractionator 168 (operating here as a depentenizer), which in oneembodiment operates at 1200 kPa with 30 theoretical stages and a refluxratio of 1.0. The temperature in the fractionator 168 typically is inthe range of 60-100° C. A C₂/C₃ overhead line 171 is removed from thefractionator 168. Overhead line 171 is split into line 172 and line 173.The contents of overhead line 172 are cooled in heat exchanger 174 andsent to a flash drum 176. The light fraction from the flash drumcomprising ethylene and propylene are purged in line 178 and can berecycled to the ethylene/propylene recovery section of an ethylenecracker. The C₄s-C₅s from the bottoms of the flash drum in line 156still contain a significant amount of butene-1 and are recycled in line156, which is combined with line 150 to form line 158, the feed line forthe autometathesis reactor. The material in overhead line 173, whenoperating in C₄ mode, is condensed in a condenser 174 to form reflux fortower 168.

The fractionator bottoms line 180, separated to 98 mol % hexene-3, isdivided into a hexene-3 line 182, a reboiler line 183, and a C₆ recycleline 188. In C₄ operating mode, the hexene-3 in line 182 fills a C₆storage tank 184 and is used as the feed for the second phase of thecampaign operation.

After production of a sufficient amount of hexene-3, the system is shutdown and prepared for operation in C₆ service. The isomerization reactor144, tower 114 and tower 168 are prepared for C₆ service. Theautometathesis reactor is not used in the second campaign. The hexene-3is fed from storage tank 184 in line 132. Line 132 becomes isomerizationfeed line 134. Material in line 134 is vaporized in heat exchanger 136and heated in heat exchanger 138, then further heated in furnace 140,and fed as line 142 first to the isomerization reactor 144. In theisomerization reactor 144, approximately 8.9% conversion to hexene-1occurs. The reactor effluent in line 147 is sent to the fractionator114. Both the isomerization reactor 144 and fractionator 114 are nowoperating in C₆ service.

To make the campaign system work, the fractionation and isomerizationequipment must be identical for C₄ and C₆ processing. The separation ofthe hexene-1 from the mixed hexene stream requires more fractionationthan the separation of butene-1 from the butene-1/butene-2 stream in C₄operation. One option is to design a tower oversized for C₄ operationbut that will be appropriate for C₆ service. Another option is to usetower 168 to provide the additional fractionation capacity for C₆operation since the autometathesis reactor is not used in C₆ operation.In this embodiment, tower 168 is used as the top portion of the C₆fractionation receiving the overhead from tower 114.

The fractionator 114 bottoms stream of C₇ ⁺ in line 118 is purged fromthe system. A side draw 122 of hexene-2 and hexene-3 is taken from a lowstage of the fractionator 114 (with the top defined as stage 1) andrecycled to the isomerization reactor 144 via lines 126 and 130. It ismixed with fresh hexene-3 feed 132 from the C₆ storage tank 184 to formisomerization reactor feed 134′ (now operating in C₆ service). Purgeline 128 is not in service. The overhead line 116 is separated to adesired ratio of hexene-1 and most of it is sent in line 148 to line 154and then line 170 to fractionator 168 (formerly functioning as adepentenizer for C₄ service). It is noted that the metathesis reactor isnot involved in the processing of the C₆ line.

In the fractionator 168, comonomer grade hexene-1 (98.5 mol %) is takenas overhead product in line 171. There is no flow through line 172 andseparator 176 is not in C₆ service. The hexene-1 product is removed inline 173, and a reflux line 175 returns to the tower with the hexene-1product being removed in line 181. The bottoms hexene-2 and hexene-3from tower 168 are recycled to the isomerization section in line 188.The contents of bottoms line 188 is mixed with the otherhexene-2/hexene-3 recycle in line 122 from tower 114 to form line 126.The fractionator 168 operates at an overhead pressure of 50 kPa.Temperature profiles for the C₆ fractionation in fractionator 114 andfractionator 168 are shown in FIGS. 5 and 6, respectively.

As indicated above, the shared equipment from the batch processflowsheet is designed for operation in both C₄ and C₆ service. Dependingon the type of equipment, this can be handled in different ways. Heatexchangers, for example, may have varying temperature approaches, butthe heat exchanger surface area may be adjusted by using multipleshells. Reactors can be addressed by using multiple reactors. Becausethe fractionator towers cannot be handled in the same way as the heatexchangers or reactors, their design is chosen to remain fixed betweenservices.

In order to use the same tower as both the C₄ and C₆ fractionator and touse the second tower as both the depentanizer and the C₆ fractionator,the tower sizing must be identical for the chosen flow rates. Becausecampaign operation allows for independent variation of the flow ratesbetween C₄ and C₆ service, operating time can be used as a variable toadjust the flow rates such that a net yearly production of, e.g., 5 KTAhexene-1 is achieved. Using this approach, in one embodiment the C₄process is operated for 2,000 hours to produce 2,696 kg/h hexene-3. 500hours of downtime is provided to empty the fractionators and reactors inpreparation for the C₆ run. Then, the C₆ process is operated for 5,333hours, feeding 1,010 kg/h hexene-3 to produce 937 kg/h hexene-1. Anadditional 500 hours is allowed for the transition back to C₄ service.Operation at these flow rates yields the same tower diameter afterchoosing 100 theoretical stages for the C₄/C₆ fractionator, and likewisefor 30 stages in the depentanizer/C₆ fractionator. An example of aprocess using the configuration of FIG. 2 is provided below as Example2.

Overall, the shared use of the fractionator and isomerization systemcomponents in the batch process eliminates 35 of the 64 pieces ofequipment from the continuous hexene-1 process. The continuous processhas 2 complete superfractionator/isomerization reactor systems comparedto just one for the campaign operation. The estimated reduction in totalinstalled capital cost is about 35-45%. This makes campaign operationespecially suited for smaller capacity installations.

Referring now to FIG. 4, an embodiment is shown in which a closed-loopheat pump is used to even further improve the efficiency of the campaignprocess depicted in FIG. 3. The overall system is designated as 200.This system uses changes in pressure to adjust the boiling point of aheat transfer fluid within the temperature ranges of a fractionator'sreboiler and condenser so that the fluid can be alternately condensedand vaporized, respectively, thus providing heat integration in place ofconventional utilities such as refrigeration for condensing or steam forreboiler. The heat pump is associated with the fractionator 214. (A topline 217 and a bottoms line 219 are removed from the fractionator 214).The fluid begins the cycle as a vapor in vapor line 202. Vapor line 202is compressed in a compressor 204 to a pressure where the temperature atwhich the vapor condenses is above the temperature of the reboiler 206.Compressed line 202 is divided into line 208 and line 210 and the twolines are recombined as line 212. Line 210 is cooled in a heat exchanger213. The use of cooling water exchanger 213 allows for control of theduty in reboiler 206 by controlling the temperature and/or flow of thehot higher pressure vapor. In some cases the contents of line 210 arecondensed in exchanger 213 and a 2 phase mixture fed to reboiler 206.This effectively limits the amount of heat transfer fluid that iscondensed in the reboiler and thus the heat transferred to the reboiler.Line 212 is condensed in the reboiler 206. Controller 236 determines thesplit between lines 208 and 210. Line 212 is now in liquid phase.

At this point in the cycle the hot duty requirement of the reboiler issatisfied. The system must now satisfy the lower temperature duty of thetower condenser. Liquid in line 212 is let down in an expansion,controlled by the overhead vapor from expansion drum 222 to drop theboiling point below the temperature of the fractionator condenser 224.This reduces the temperature of the line as a portion of 212 isvaporized at the lower pressure. Line 212 is then combined with(optional) line 215 as line 216. Optionally, as described below, line216 is cooled in heat exchanger 218 using cooling water line 220. Vaporis removed from drum 222 as line 228. Vapor is not sent to the condenser224. The liquid is removed from drum 222 in line 226. Line 226 is thenvaporized in the condenser 224 and after vaporization is combined withline 228 to form line 230. Line 230 passes through a knockout drum 232.The vapor in drum 232 is removed as line 202, which as mentioned above,is returned back to the compressor 204. Any small amount of liquid indrum 232 is removed in line 215, which is combined with the effluentfrom the expander 240 to form line 216.

To settle any difference in condenser and reboiler duty, an additionalexchanger may be required within the loop to add or remove heat asnecessary. In this particular campaign system, the condenser duty isgreater than that of the reboiler, so heat exchanger 213 is placed atthe compressor discharge in parallel with the reboiler to remove thedifference in heat duty, allowing the heat pump fluid to fully condensein the reboiler 206. Heat exchanger 213 is the exchanger to remove“extra” duty from the loop based on the reboiler-condenser dutydifference. Conversely, if the reboiler duty were greater, a heatexchanger 238 operated with steam instead of cooling water would beplaced at the outlet of the expander 240 to have more of the fluid asvapor to separator 222 and thus feed lower amount of liquid as line 226to vaporize in the condenser.

To avoid temperature crosses in the reboiler and condenser, atemperature difference of, e.g., 3° C. can be used between the outlettemperatures of the process fluid, line 219 for example, and the heatpump, line 212. The compressor and expander discharge pressures can beconsidered fixed, set by the requirement to bring the heat pump fluidboiling point within the approach to the reboiler and condensertemperatures, respectively. Effectively, the work performed by thecompressor is that required to undo the work of the expander, so theenergy costs of the heat pump decrease as the temperature spread acrossthe tower becomes smaller.

With the pressures in each portion of the heat pump cycle fixed and thecondenser as the controlling duty, the inlet temperature to thecondenser can be maximized in order to provide additional coolingcapacity. In this case the cooling water heat exchanger 218 is used inthe loop at the outlet of the expander to cool the line before enteringthe condenser 224. This is a low-cost method of reducing the circulationrate through the higher-cost compressor 204.

The working fluid for the heat pump used in a campaign mode ofprocessing typically is a hydrocarbon or mixture of hydrocarbons suchthat the boiling point of that hydrocarbon or mixture falls between theboiling point of the first carbon number and the second carbon number.This differs from a conventional closed loop heat pump system that isoperating on a single carbon number. There, the working fluid isselected based upon a single carbon number and typically has propertiesclose to the hydrocarbon being separated. In one particular closed-loopsystem for campaign mode operation, n-butane is used as the circulatingfluid and the heat pump is applied to one or both of the fractionatorsin the isomerization section operating in either C₄ or C₆ service.N-butane boils between the 1-butene overhead in one mode and the3-hexene reboiler in the C6 service mode. Note also it is possible touse a mixture of fluids as the heat transfer fluid. In that case it ispossible to adjust the composition to optimize the thermodynamicproperties of the fluid mixture. The energy requirement of the heat pumpis that to run the compressor, significantly less than the stand-alonetower due to the integration of reboiler and condenser heat duties. Aclosed loop applied to both towers would integrate both reboilers inseries within the loop, then both condensers in series, adjusting thecompressor and expander discharge pressures accordingly.

Variations in process parameters equipment sizing, etc. will depend uponthe desired butene-1 purity. Very high purity butene-1 generally is notrequired to produce hexene-3 by autometathesis. The fractionation towercould be designed to produce for example 95% butene-1, with one part ofthe product going to autometathesis for lights, recycle and hexene-3,and another part going to a different fractionator to produce highpurity butene-1 (polymer grade). In this case, both the depentanizer ofthe main example and the additional butene-1 fractionation could be usedin campaign mode for hexene purification. It is noted that feed toautometathesis could be a tower side draw where overhead is the highpurity butene-1.

A number of other methods of isomerization and metathesis can be appliedand still maintain the important feature of the campaign process, whichis shared equipment between C₄ and C₆ service. Further, additionalprocessing steps and/or alternate feedstocks can be used. Onenon-limiting example of an alternative process would be to employ anadditional metathesis reaction step involving the reaction of ethylenewith butenes (ethenolysis). Depending upon the feed quality to theautometathesis step (the butene-2 content) some propylene and pentene-2will be formed. A second autometathesis step involving the reactionbetween butene-1 and pentene-2 to yield propylene and hexene-3 could beincluded. This additional autometathesis step would, for example,involve line 75 or 82 of FIG. 2. Pentene-2 produced in reactor 52 can besent to a second autometathesis reactor to make more hexene-3. A furtherexample would be to incorporate the process of U.S. Pat. No. 4,709,115(Jung et al, November 1987), where the metathesis step occurs in acatalytic distillation tower. In the campaign system, this variationwould manifest itself as a catalytic depentanizer in C₄ service, withthe catalyst replaced by inert beads to operate as C₆ fractionator in C₆service.

The campaign processing scheme described herein also can be integratedwith other processing units. For example, the campaign process can beintegrated with a continuous conventional metathesis unit for producingpropylene from the reaction of ethylene and butene-2. The conventionalmetathesis process typically feeds ethylene and a mixed C₄ raffinatestream containing butene-2, which can be the same C₄ feed raffinatestream as that used for the process in FIG. 1, to a metathesis reactor.Ethylene and propylene produced in the autometathesis step of thecampaign operation is then sent to the fractionation system of theconventional metathesis unit. This effectively provides feed ethylene tothe conventional metathesis reaction and provides propylene product. Arecycle of C₄ is taken from the second separation step back to thereactor. To integrate the campaign process in FIG. 3, for example, thepurge streams 118, 126, and 178 would be recycled to an appropriatedisposition in the conventional metathesis process. Because thisintegration can use the same C₄ raffinate feed stream, a number ofoptions are afforded. One option is that the total C₄ raffinate canremain constant, thus converting some propylene production inconventional metathesis to hexene-1 production in the campaign process.A second option is to maintain constant propylene production from theconventional metathesis process and “scale up” up the C₄ raffinate flowrate as required for butene-1 or hexene-1 production from the campaignprocess. Any intermediate combination of product flow rates from theseoptions is also possible.

Two significant advantages exist in integrating the campaign processwith conventional metathesis. First, all purge streams from the campaignprocess which would otherwise be lost to the cracking furnace or anotherlow-value disposition can be recovered in the conventional process. Inparticular, the C2-C3 of stream 78 is ethylene that has been upgradedfrom the C₄ raffinate feed and greatly reduces the fresh ethylenerequirement to the conventional metathesis unit. Second, it is possibleto make use of the autometathesis reactor of the campaign process whenit is idle in C₆ service. If additional C₄ raffinate and ethylene areavailable, they can be fed to the autometathesis reactor during the C₆phase of the campaign process. Metathesis will produce additionalpropylene with no change of catalyst required. The additional propylenecan be fed to the separation equipment of the conventional metathesisprocess and recovered. In this way, extra propylene production for aportion of the operating year is possible either to the limit of theseparation equipment overdesign in a retrofit case, or to a desiredamount in the case of a new integrated plant.

In a second integration example, a process for the production of olefinsfrom methanol (MTO) produces a C₄-C₆ stream composed of linear olefinscan be incorporated upstream. Similarly, processes that utilize ethyleneoligomerization to produce linear alpha olefins have alpha olefinstreams in even carbon numbers from C₄ through C₂₀+ can be incorporated.The campaign processing scheme in combination with metathesis can beused to adjust the carbon number distribution and thus maximize productvalue dependent upon market conditions. For example, a C₁₀ alpha olefincan be isomerized to a number of C₁₀ internal olefins using theisomerization step. The internal olefins can be then reacted withethylene in a metathesis step to produce a range of lower carbon numberalpha olefins. These can be separated and processed in campaign mode oralternately the isomerization/metathesis process used for a C₁₆ alphaolefin following operation with a C₁₀ alpha olefin feed.

The particular campaign system described includes two fractionators. Oneis used as a C₄ and C₆ fractionator and combined with the isomerizationreactor. The second is used as a depentanizer and second C₆fractionator. In one embodiment, the heat pump is placed on only theC₄/C₆ fractionator, with the depentanizer/C₆ fractionator usingconventional utilities. Another embodiment involves use of a heat pumpon the depentanizer/C₆ fractionator. In a further configuration a singleheat pump loop passes through all fractionation towers. For a campaignprocess, this type of loop would operate on the C₄ fractionator anddepentanizer during C₄ service, then on both C₆ fractionators during C₆service, with possible sequences ofreboiler-reboiler-condenser-condenser orreboiler-condenser-reboiler-condenser within the loop. Yet anotherconfiguration places one heat pump loop on each tower.

The embodiment of FIG. 2 shows a two-tower system for the fractionatorfollowing the isomerization system. This embodiment also can include theoption of an integrated condenser/reboiler, with a heat pump on the“outer” exchangers. In this configuration, the tower pressures would beadjusted such that, for example, the condenser of the C₄ fractionatorcould be integrated with the reboiler of the depentanizer. The heat pumpis then placed across the wider temperature range of the remainingexchangers. The integrated condenser/reboiler is possible with eitherthe upstream or the downstream tower operating as the higher pressuretower.

In the embodiments described above in detail, one suitable circulatingfluid to be used in the heat pump is n-butane. This fluid is usefulbecause its boiling temperature is within range of the tower condenserand reboiler at appreciable pressure. Alternative heat pumps on thistype of system can use a mixture of hydrocarbons or other fluids.Mixtures are particularly useful to extend the boiling range of the heatpump fluid for towers with wide temperature profiles, thus minimizingthe difference between compressor and expander discharge pressures.

Heat pump loops with a constant circulation rate between C₄ and C₆service are possible, with the compressor and/or expander dischargepressures relaxed to allow the use of some sensible heat when in thelower-duty service. A heat pump loop with or without heat exchanger 118,placed to provide additional cooling capacity at low cost (coolingwater), is also feasible. In lieu of balancing the condenser andreboiler duties by removing heat in heat exchanger 113, a steamexchanger could be placed to add heat in the opposite portion of theloop.

The heat pump can be expanded to both fractionators of the batch system.This would result in heat integration of the condensers and reboilers ofthe C₄ fractionator and depentanizer, and the two C₆ fractionators.There can be a single heat pump loop or two separate loops having oneheat pump on each tower.

One or more of the above variations may be used with an open-loop heatpump. An open-loop heat pump uses the tower overhead stream as the heatexchange fluid against the bottoms stream.

In another embodiment the pressures of the two fractionators can beadjusted such that, for example, the condenser of the C₄ fractionator isheat-integrated with the reboiler of the depentanizer. A heat pump canthen be used on the reboiler of the C₄ fractionator and condenser of thedepentenizer.

Some or all of the heat pump configurations described above may beemployed using a mixture of fluids in the heat pump loop, which wouldprovide a broader region of vaporization/condensing than a pure fluid.The required difference between compressor and expander dischargepressures, thus the compressor energy consumption, would be decreased.The batch system also presents the possibility of different fluids inthe heat pump loop between C₄ and C₆ service, or a different fluid foreach loop if two loops are used.

The following examples are included to illustrate certain features ofthe disclosed embodiments but are not intended to limit the scope of thedescription.

Example 1 Campaign Autometathesis Process Employing High-Selectivity W03Catalyst and an Integrated Two Tower Fractionation System

A computerized simulation was conducted in which butene isomerizationand autometathesis sections are operated as one unit, temporarilystoring n-C₆ made in the autometathesis section in a process having theconfiguration shown in FIG. 2 with storage tank 84 being used but not41. This was followed by C₆ isomerization operation utilizing theequipment used in the butene isomerization section. In this scheme, theentire set of C₆ isomerization equipment is avoided to reduce capitalcost. In this Example, 5 KTA of polymer grade 1-hexene was produced in acampaign process. The butene feed used in shown in table 1 below.

TABLE 1 C4 Feed to Autometathesis Process Component Wt % iso-butane 4.0n-butane 18.1 tr2-butene 18.2 1-butene 50.5 iso-butene 0.10 cis2-butene11.1 Total 100.0 Flow Rate, Kg/H 14,900

In the campaign operation simulation, C₄ isomerization andautometathesis sections were operated for 2000 hours producing 3-hexene,which was temporarily stored in a storage tank for further isomerizationto 1-hexene. After producing the 3-hexene for 2000 hours, C₄isomerization and autometathesis operation were shut down and thedistillation towers and reactors were emptied. The C₆ isomerizationsection was then operated for 5333 hours producing 1-hexene from thestored 3-hexene, utilizing the same equipment used in C₄ isomerizationoperation. Eliminating the C₆ isomerization equipment reduces capitalcost. The particular hours of operations were chosen such that theinterchangeability of the equipment was possible providing for a netyearly production of 5000 KTA of 1-hexene.

This process followed the same scheme as would be used in a continuousautometathesis process. The separation of 1-hexene from its isomers wasthe critical separation. In the continuous process, a two-tower designis used for this separation. The same two-tower separation system wasused in the batch process for separation of 1-hexene from other C₆compounds. Since this equipment was used in the butene isomerization,the same two-tower separation system applies to separation of 1-butenefrom 2-butene as well.

The raffinate II feed composition is given in Table 1. The 1-butenecontent in the C₄ feed is higher than the equilibrium butene ratio at650 F, the operating temperature of the C₄ isomerization reactor (feedB1/B2=2.8, equilibrium B1/B2 at 650 F=028). Hence the raffinate II feedwas sent to the C₄ separation tower system to separate 1-butene prior toentering the butene isomerization reactor.

The bottom product from the C₄ separation contained mainly 2-butenes andn-butane. This bottom product stream was recycled to the isomerizationreactor to increase n-butene utilization. A small purge was removed fromthis recycle stream to control the build-up of the inerts, n-butane andiso-butane. The isomerization reactor feed exchanged heat with the hotreactor product. The reactor feed was further heated to the reactiontemperature inside a fuel-fired furnace and entered the reactor. Thereactor was operated at 650 Deg. F and 117 psia. The catalyst was MgOtablets. Feed ratios and 2-butene conversion data for the isomerizationreactor are shown below on Table 2. The reaction product was sent to thebutene separation system. The C₄ isomerization reactor and separationtower were used during the C₆ isomerization operation as well

The C4 separation system consisted of a two tower system. The condenserof one tower is used to reboil the second tower. The two towers areoperated at different pressures to allow for this exchange. Splittingthe feed with a portion to each tower reduces energy consumption whilebalancing the duties for each tower. The first fractionator had 80stages and second fractionator had 70 stages. The raffinate II feedentered second tower at stage 24. The butene isomerization reactorproduct was split. One portion entered the first tower at stage 15 andother portion entered the second tower at stage 48. The distillateproduct from first tower, concentrated in 1-butene entered the secondtower at stage 30. By adjusting the vapor feed split ratio and operatingpressures of the towers, energy exchange between the tower 1 condenserand tower 2 reboiler was made possible. The final distillate product was90 mol % 1-butene, which was sent to the autometathesis section forfurther processing. The 1-butene product stream contained iso-butane(5.1 wt %), n-butane (3.8 wt %), tr2-butene (1.2 wt %) and iso-butene(0.13 wt %). If required, monomer grade 1-butene (99 wt %) could also beproduced from this separation system. Details of the separation towerare given in Table 5 below.

In the second processing step but still operating in the C4 mode,1-butene from the butene isomerization/separation system was sent to theautometathesis section to produce n-hexenes. In this section, 1-butenefeed was mixed with recycled 1-butene from the separators and exchangedheat with the hot reactor product. The reactor feed was further heatedto the reaction temperature inside a fuel-fired furnace and entered thereactor. The autometathesis reactor operated at 600 F and 275 psia. Thecatalyst was WO3 on high-purity silica. Inside this reactor, 1-butenereacted with itself to produce ethylene and 3-hexene. Side reactionsbetween 1-butene and 2-butenes producing propylene and 2-pentenes alsooccurred. Moreover, isobutylene reacted with 1-butene to produceethylene and 2-methyl-2-pentene (i-C₆ olefin, BP=67.3 deg C). The otherpossible isobutylene reactions were determined to be insignificant. Theproduction of i-C₆ olefin in the autometathesis reactor is undesirablesince it affects the purity of 1-hexene product. Hence the isobutylenecontent in the raffinate II feed was kept very low. A small amount of C7and C8 was also produced by other metathesis reactions. Recycle,conversion, and reaction product composition data are shown on Table 3.

The autometathesis reaction products were separated in the depentenizer.The 3-hexene was recovered in the depentenizer tower as bottom productand sent to the storage tank for use in the C6 mode step of the campaignprocess. The i-C₆ olefins, C7 and C8 produced inside the autometathesisreactor were also carried along with the 3-hexene. The distillate fromthe depentanizer was sent to the depropenizer for further separation.The lighter components, ethylene and propylene were recovered asdistillate and sent to product recovery in the ethylene plant. Theunconverted 1-butene was recovered as the bottom product and recycled tothe autometathesis reactor to improve butene utilization. A side-drawstream was purged from the depropylenizer to remove the inerts,iso-butane and n-butane from the autometathesis system. The details ofthe separation tower are shown below.

At the completion of the C₄ isomerization and autometathesis run, theprocess was shut down. The reactors and distillation towers were emptiedin preparation for the C₆ isomerization run. The C₆ isomerizationoperation was then conducted as shown in the process flow diagram.

The hexene isomerization section consisted of a hexene isomerizationreactor and hexene separation system, and the same equipment as was usedfor C₄ processing was employed. The 3-hexene from the storage tank wasmixed with recycled 2-hexenes and 3-hexenes from the C₆ separationsystem and exchanged heat with the hot isomerization reactor product.The isomerization reactor feed was further heated to the reactiontemperature inside a fuel-fired furnace and entered the reactor. Thereactor operated at 650 F and 56 psia. The catalyst was MgO tablets. Thereactor product was an equilibrium mixture of 1-hexene, 2-hexenes and3-hexenes including the cis-trans isomers. This product mixture wasseparated in the hexene separation system to produce polymer grade1-hexene as distillate product. (The two-tower separation system withenergy integration was explained previously.) The bottom product,2-hexenes and 3-hexenes were recycled to the isomerization reactor. Asmall purge was taken from the separation system to remove the heavycomponents from the C₆ isomerization system. The details of theseparation tower are given below in Table 6.

The 2-methyl-2-pentene from the autometathesis reactor was also carriedover to the hexene isomerization section where the isomerizationactivity of the MgO catalyst produced its isomers: 2-methyl-1-pentene,4-methyl-1-pentene, 4-methyl-cis-2-pentene and 4-methyl-trans-2-pentene.Since the boiling point of all i-C₆ except the 2-methyl-2-pentene arelower than 1-hexene, any i-C₆ produced in the autometathesis reactorends up with the 1-hexene product.

TABLE 2 Butene Isomerization Reactor Rx Operating Temp, F. 650 RxOperating Pr, Psia 117 Catalyst MgO tablets Rx Feed B2/B1 49 Rx ProdB2/B1 ratio 3.6 2-Butene conversion % 21.7

TABLE 3 Autometathesis Reactor Rx Operating Temp, F. 600 Rx OperatingPr, Psia 275 Catalyst WO3 on high purity silica Rx Feed B1/B2 ratio 96.2B1 Conversion mol, % 46.2 Molar Selectivity, % Ethylene 40.53 Propylene12.22 Pentene 0.02 n-hexene 46.48 i-hexene 0.15 C7 and CB 0.60

The autometathesis reactor performance was based on experimental datafor high selectivity WO3 catalyst. This information was incorporatedinto the HYSYS simulation. The conversion and selectivity weredetermined for the autometathesis reactor feed given in Table 10. Theautometathesis selectivity for the main reaction (C₂+C₆) was 87.01. Theselectivity for the side reactions (C₃+C₅) was 12.37. The selectivityfor the isobutylene reaction with 1-butene was 0.15. A small amount ofC7 and C8 was also formed in the autometathesis reactor.

TABLE 4 Hexene Isomerization Reactor Rx Operating Temp, F. 650 RxOperating Pr, Psia 56 Catalyst MgO tablets Rx Feed (2-hex + 3-hex)/1-hexratio 63.36 Rx Prod (2-hex + 3-hex)/1-hex ratio 10.6 1-Hex ProdComposition, mol % 8.3

The C₆ isomerization reactor performance was obtained from thecorrelation of experimental data. This correlation was incorporated intothe HYSYS simulation.

TABLE 5 Specifications of the Separation Columns in C4 Isomerization andAutometathesis Butane Butane Parameter Depropartizer DepornenizerSplitter1 Splitter2 * Number of Stages 15 40 80 70 Feed Tray (# from 520 15 24, 30, 48 top) condenser P, Kpa 2200 1600 700 530 re-boiler P,Kpa 2300 1800 750 550 Top Spec 1.5 mol % 1-C4 in 0.1 mol % n-C₆ 40 mol %1- 90 mol % 1- distillate in distillate butene in butene in distillatedistillate Bottom Spec 0.5% propylene in 0.01 mol % n-C₅ 0.5 mol % 1 mol% 1- bottom in bottom 1-butene in butene in product bottom bottom prodproduct Other Specs Top vent = 15 kmol/h Side draw = 23 kmol/hInterstage cooling at stage 3 = 1000 KW Note Bottom product is 98.5 wt %n-C₆

TABLE 6 Specifications of the separation columns in C₆ isomerization C₆C₆ Parameter Splitter1 Sp1itter2 * Number of Stages 80 70 Feed Tray (#from 25 40.60 top) condenser P, Kpa 207 117 re-boiler P, Kpa 241 138 TopSpec 60 mol % 2 & 3 1.2 mol % 2 & 3 hexene in hexene in distillatedistillate Bottom Spec 1 mol % 1-hexene 2.5 mol % 1- in side-draw hexenein bottom product prod Other Specs 45 mol % C7 & C8 in bottom productNote Side-draw from stage 74

The material balance for the batch case, producing 5 KTA of polymergrade 1-hexene is given below. The material balance summary as wellcompositions of key streams are given in the following tables.

TABLE 7 Overall material balance for C4 Isomerization and Autometathesis(2000 hours operation) MTA Feed C4 Feed 12,000 Total Feed 12,000Products C2/C3 to cracker  1,852 Depropenizer side draw  2,120 C4 purge 2,637 Depentenizer Bottom 5,391 Total Products 12,000

TABLE 8 Overall material balance for C₆ Isomerization (5333 hoursoperation) MTA Feed C₆ Feed 5,391 Total Feed 5,391 Products Hexane-1Prod 5,026 C₆ purge 205 C₆ + Purge 160 Total Products 5,391

TABLE 9 Material balance for the butene isomerization section (2000hours operation) C4 C4 Component, C4 C4 Isomerization Isomerization C4B1 to wt % Feed recycle Feed Prod Purge Autometathesis Iso-butane 4.040.0 0.0 0.0 0.0 5.2 n-butane 16.14 61.3 61.3 61.3 61.3 3.5 Tr2-butene18.17 22.6 22.6 17.8 22.6 1.3 1-butene 50.45 0.70 0.70 8.5 0.70 89.7Iso-butene 0.10 0.0 0.0 0.0 0.0 0.13 Cis2-butene 11.10 15.4 15.4 12.315.4 0.10 Total, wt % 100 100 100 100 100 100 Flow, Kg/h 6000 1516515165 15165 1318 4682

TABLE 10 Material balance for the autometathesis section (2000 hoursoperation) Auto. Auto. C2/C3 Depro- Depen- Component, B1 to B1 Rctr.Rctr. to penizer tenizar wt % Autometathesis Recycle Feed Prod Crackersidedraw bottom Ethylene 0.0 0 0 6.0 62.4 17.1 PropyleNe 0.0 0.4 0.2 2.916.7 17.1 Iso-butane 5.2 20.0 14.6 14.6 6.3 18.0 n-butane 3.5 26.1 17.917.9 3.7 12.7 Tr2-butene 1.3 0.2 0.6 0.2 0 0.10 1-butene 89.7 48.0 63.334.0 10.9 34.5 Iso-butene 0.13 0.04 0.07 0.03 0 0.04 Cls2-butene 0.100.04 0.06 0.02 0 0.0 n-Pentene 0.0 5.1 3.2 3.2 0 0.41 3-hexene 0.0 0.20.13 20.8 0 0.0 98.2 i-C₆ 0.0 0.0 0.0 0.07 0 0.0 0.3 C7 & C8 0.0 0.0 0.00.32 0 0.0 1.5 Total, wt % 100 100 100 100 100 100 100 Flow, Kg/h 46828,096 12,799 12,799 9.26 1,060 2696

TABLE 11 Material balance for the hexens isomerization section (5333.3hours operation) C₆ C₆ isomer- isomer- 1- Component, C₆ ization izationhexene C₆ C_(r) + wt % Recycle Feed Prod Prod Purge. Purge 1-hexene 1.71.6 8.4 98.5 1.7 0.2 Tr2-hexene 46.1 50.0 42.9 0.46 46.1 23.6 Tr3-hexene21.5 19.9 20.0 0.35 21.5 8.6 Cis2-hexene 23.1 21.4 21.5 0 23.1 16.3Cis3-hexene 7.1 6.6 6.6 0.4 7.1 2.4 i-C₆ 0.1 0.1 0.1 0.33 0.1 0.1 C7 &C8 0.44 0.5 0.5 0.0 0.44 48.8 Total, wt % 100 100 100 100 100 100 Flow,Kg/h 12,645 13,656 13,656 942 38 30

A comparison of the overall material balances for a continuous processat 50 KTA using the same feed composition as the batch process of thisexample shows that the major streams scale down linearly for thecampaign case. Some minor difference in the C2/C3 to cracker anddepropenizer side-draw were noticed, and arose from the operation of thedepropenizer tower. The energy balance for the campaign case study,producing 5 KTA of polymer grade 1-hexene is provided below. The energybalance for the butene isomerization section, autometathesis section andhexane isomerization sections are shown. In Table 12 there are shown twobalances. The “Before Exchange” tabulation lists the duties for each ofthe tower reboilers or condensers in the C4 or C6 modes. The “AfterExchange” tabulation simply subtracts the common duty from the “beforeExchange” tabulation. For example in the C4 Isom mode, Tower 1 condenserhas a duty of 6733 KW and the tower 2 reboiler has a duty of 6670 KW.Since these are exchanged against each other, the after exchange duty isthe difference (63 KW).

TABLE 12 Energy balance summary for the campaign process C₄ isom C₆ isom(2000 Autometathesis (5333 hours) (2000 hours) hours) Total BEFOREEXCHANGE Feed vaporizer (LPS), KW 1626 980 1408 Feed Heater (fuel), KW130 430 126 Tower 1 Condenser duty 6733 (a) © 1000, 71 * 5876 (a) (CW),KW Tower 1 Re-boiler duty 5437 1500 4801 (LPS), KW Tower 2 Condenserduty 7843 2959 6414 (CW), KW Tower 2 Re-boiler duty 6670 (a) 1427 **5853 (a) (LPS), KW Pump, power, KW 98 50 90 AFTER EXCHANGE Feedvaporizer (LPS), KW 1626 980 1408 Feed Heater (fuel), KW 130 430 126Tower 1 Condenser duty 63 (a) © 1000, 71 * 5876 (a) (CW), KW Tower 1Re-boiler duty 5437 1500 4801 (LPS), KW Tower 2 Condenser duty 7843 29596414 (CW), KW Tower 2 Re-boiler duty 0 (a) 1427 ** 0 (a) (LPS), KW Pump,power, KW 98 50 90 Total Utility After Exchange FUEL, KW 130 430 126 −5REF, KW 71 CW, KW 7906 3959 6437 LPS, KW 7063 2480 6209 HPS, KW 1427POWER, KW 98 50 90 Total Utility After Exchange (8000 Hour Basis) FUEL,KW 224 −5 REF, KW 17.8 CW, KW 7,258 LPS, KW 6,525 HPS, KW 357 POWER, KW97 FUEL, MKCAL/H 0.19 −5 REF, MKCAL/H 0.015 CW, MKCAL/H 6.22 LPS,MKCAL/H 5.59 HPS, MKCAL/H 0.306 POWER MKCAL/H 0.083 Note: 1. InAutometathesis, towerl is a depropylenizer and tower 2 is adepentenizer. (i) * - depropenizer condenser is - 5 deg refrigerant,(ii) ** - depentenizer re-boiler is HPS. (iii) © 1000 KW CW interchangeon DEC3 to reduce the refrigerant duty. 2. In C₄ Isom, towerl is BS1 -higher pressure and tower 2 is BS2 - lower pressure tower. 3. In C₆Isom, towerl is HS1 - higher pressure and tower 2 is HS2 - lowerpressure tower. 4. The energy integration in two-tower system wasexplained previously. a. exchange between BS1/HS1 condenser and BS2/HS2re-boiler in C₄C₆ isomerization system. This is the exchange for theinternal condenser/reboiler system of the two tower split feed system

The energy balance for the batch process given in Table 12 above can becompared to the energy balance for a continuous process using the sameequipment and feed composition, and is shown below on Table 13.

TABLE 13 Energy balance summary for a continuous processingautometathesis case before and after energy exchange C₄ Isom Automet C₅Isom Total Feed vaporizer (LPS), Kw 4141 2447 9747 Feed Heater (fuel),KW 971 1088 1816 Tower 1 Condenser duty (c) 2500, 80 * 24920 (a) (CW),KW Tower 1 Re-boiler duty 3756 19810 (LPS), KW Tower 2 Condenser duty35110 7551 31070 (b) (CW), KW Tower 2 Re-boiler duty 28400 (b) 3682 **23980 (a) (LPS), KW Pump, power, KW 104 50 165 Total Utility BeforeEnergy Exchange FUEL KW 3,875 −5 REF, KW 80 CW, KW 101,151 LPS, KW92,281 HPS, KW 3,682 POWER, KW 319 FUEL MKCAL/H 3.32 −5 REF, MKCAL/H0.07 CW, MKCAL/H 86.7 LPS,MKCAL/H 79.10 HPS, MKCAL/H 3.16 POWER, MKCAL/H0.27 Total Utility After Energy Exchange FUEL, KW 3.875 −5 REF, KW 80CW, KW 48,771 LPS, KW 39,901 HPS, KW 3,682 POWER, KW 319 FUEL, MKCAL/H3.32 −5 REF, MKCAL/H 0.07 CW, MKCAL/H 41.80 LPS, MKCAL/H 34.20 HPS,MKCAL/H 3.16 POWER, MKCAL/H 0.27 Note: 1 . In Automet, towerl is adepropenizer DeC3) and tower 2 is a depentenizer (DeC5). (i) * -depropenizer condenser is - 5 deg refrigerant, (ii) ** - DEC5 re-boileris HPS. (iii) © 2500 KW CW interchange on DEC3 to reduce the refrigerantduty. 2. In C₆ Isom, tower1 is HS1 - higher pressure and tower 2 isHS2 - lower pressure tower. 3. The energy integration was explained intable.17. It is noted in table 22 as well. (a) exchange between HS1condenser and HS2 re-boiler in C₆ isom system. (b) exchange between HS2condenser in C₆ isom and BS re-boiler in C₄ Isom. It is noted thatenergy integration reduced the total cooling water requirement by 53%and LPS requirement by 56%.

The utility usage scales down linearly in the batch case except for thecooling water and LPS. In the batch operation, an energy efficienttwo-tower C₆ separation system is employed. However, the energyintegration between C₆ isomerization and C₄ isomerization towers couldnot practiced due to the batch operation. This increased the CW and LPSusage in the batch process. It appears that the savings in capital costmore than offset the increased utility cost. Details can be found in theeconomic evaluation of the processes.

Example 2 Campaign Autometathesis Process

In the previous Example, a campaign process to produce 5 KTA of 1-hexenewas discussed. In this process, the improvements over Example 1 are:

-   -   1. In the C₄ isomerization process, one distillation tower with        100 stages replaced two-tower system with total of 150 stages.    -   2. In the autometathesis section, the depropylenizer tower was        replaced by a gas-liquid separator.    -   3. In the C₆ isomerization process, one distillation tower with        100 stages (same used in C₄ isom) replaced two-tower system with        total of 150 stages. The depentanizer tower acted as the second        distillation tower.

Elimination of the two-tower separation system impacted the energy usagein the process. Since energy integration was not done, the utilityconsumption increased. Furthermore, the elimination of depropylenizertower resulted in increased purge flows. However, economic analysisshowed that capital cost savings more than offset the increased utilitycost. In this case study, 5 KTA of polymer grade 1-hexene was producedin an improved campaign process.

The butene feed used in this study is given in Table 1 above. Theprocess flow scheme is shown in FIG. 3.

The raffinate II feed was sent to the C₄ separation tower to separate1-butene prior to the butene isomerization reactor. A single-towerseparation replaced the energy integrated two-tower system in FIG. 2 inorder to eliminate one distillation tower and associated equipment.

In the autometathesis section, a gas-liquid separator replaced thedepropylenizer tower. The depentenizer vapor distillate was cooled andsent to a gas-liquid separator. The lighter components, ethylene andpropylene were recovered as vapor and sent to product recovery in theethylene plant. The unconverted 1-butene was recovered in the liquidproduct and recycled to the autometathesis reactor to improve buteneutilization. Most of the autometathesis equipment was used only duringautometathesis operation, except for the depentanizer tower that wasused in C₆ isomerization operation as well.

At the completion of C₄ isomerization and autometathesis run, theprocess was shut down. The reactors and distillation towers were emptiedin preparation for the C₆ isomerization run. The C₆ isomerizationoperation was conducted employing the same equipments as shown in FIG.3.

In the C₆ isomerization section, the feed mixture was separated usingtwo distillation towers as shown in the flow diagram. The firstdistillation tower (butene splitter) produced 93% hexene-1 distillateproduct. The bottom product, 2-hexenes and 3-hexenes were recycled tothe isomerization reactor. A small bottom purge was taken from thisdistillation tower to remove the heavy components, C₇ and C₈ from the C₆isomerization system. The depentenizer acted as the second hexenedistillation tower that produced polymer grade hexene-1 from 93%hexene-1 feed. The specification for the depentenizer was 65 mol %1-hexene in the bottom product. This specification was relaxed so that2&3-hexenes were carried to bottom product allowing polymer grade1-hexene to be made as the distillate product. The flow rate of thisproduct was low as compared to the distillate product. The depentenizerbottom was mixed with the other bottom product and recycled to the C₆isomerization reactor. By employing the depentenizer as the secondhexene splitter, the number of stages in the first hexene splitter wasreduced from 150 to 100. The details are given in following tables.

TABLE 14 Butene Isomerization Reactor Rx. Operating Temp, F. 650 RxOperating Pr, Psia 117 Catalyst MgO tablets Rx Feed B2/B1 ratio 203 RxProd B2/B1 ratio 3.6 2-Butene conversion, % 21.7

TABLE 15 Autometatbesis Reactor Rx Operating Temp, F. 600 Rx OperatingPr, Psia 275 Catalyst WOP3 On high purity silica Rx Feed B1/B2 ratio33.6 B1 Conversion, mol % 36.2 Molar Selectivity, % Ethylene 40.0Propylene 12.50 Pentene 0.83 n-hexene 45.7 i-hexene 0.15 C7 and C8 0.82

The autometathesis reactor performance was based on experimental datafor high selectivity WO3 catalyst. This information was incorporatedinto the HYSYS simulation. The conversion and selectivity weredetermined for the autometathesis reactor feed given in Table 21. Theautometathesis selectivity for the main reaction (C₂+C₆) was 85.7. Theselectivity for the side reactions (C₃+C₅) was 13.3. The selectivity forthe isobutylene reaction with 1-butene was 0,15. A small amount of C7and C8 were also formed in the autometathesis reactor.

TABLE 16 Hexene Isomerization Reactor Rx Operating Temp, F. 650 RxOperating Pr, Psia 56 Catalyst MgO tablets Rx Feed (2-hex + 3-hex)/1-hexratio 43.4 Rx Prod (2-hex + 3-hex)/1-hex ratio 10.9 1-Hex ProdComposition. mol % 8.4

TABLE 17 Specifications of the separation columns Butene Hexene HexeneParameter Splitter Depentenizer Splitter1 Splitter2 * Number of Stages100 30 100 30 Feed Tray (# from 20, 45 10 65 10 top) condenser P, Kpa570 1200 120 110 re-boiler P, Kpa 600 1280 140 120 Top Spec 90 mol % RR= 1.0 7 mol % 1.2 mol % 1-C₄ in 2&3-hex 2&3-hex distillate in indistillate distillate Bottom Spec 0.2% 1-C₄ 0.98 mol % 1.5 mol % 65 mol% in bottom C₆ in bottom 1-C₆ in 1-hex in side-draw bottom prod OtherSpecs 45 mol % C7 and C8 in bottom product Note * Note: During the C₆Isomerization batch operation, the depentenizer tower acts as the secondhexene splitter, allowing reduced number of stages on the firstseparation tower.

The material balance for the batch case study, producing 5 KTA ofpolymer grade 1-hexene is given below. The material balance summary aswell compositions of key streams are given in the following tables.

TABLE 18 Overall material balance for C₄ Isomerization andAutometathesis (2000 hours operation) MTA Feed C₄ Feed 12,250 Total Feed12,250 Products C2/C3 to cracker  4,240 C₄ purge  2,617 DepentenizerBottom  5,393 Total Products 12,250

TABLE 19 Overall material balance for C₆ Isomerization (5333 hoursoperation) MTA Feed C₆ Feed 5,387 Total Feed 5,387 Products Hexane-1Prod 4,997 C₆ purge   225 C₆₊ Purge   165 Total Products 5,387

TABLE 20 Material balance for the butene isomerization section (2000hours operation) C₄ C₄ Isomer- Isomer- B1 to Component, C₄ C₄ izationization C₄ Automet- wt % Feed recycle Feed Prod Purge athesis Iso-butane4.04 0.0 0.0 0.0 0.0 5.1 n-butane 16.14 61.8 61.8 61.8 61.8 3.7Tr2-butene 18.17 22.6 22.6 17.6 22.6 1.2 1-butene 50.45 0.20 0.20 8.40.20 89.7 Iso-butene 0.10 0.0 0.0 0.0 0.0 0.13 Cis2-butene 11.10 15.315.3 12.2 15.3 0.10 Total, wt % 100 100 100 100 100 100 Flow, Kg/h 612515029 15029 15029 1308 4817

TABLE 21 Material balance for the autometathesis section (2000 hoursoperation) Component, Butene-1 to Butene-1 AR AR C2/C3 to DepropenizerDepentenizer wt % Autometathesis Recycle Feed Prod Cracker sidedrawbottom Ethylene 0.0 5.7 4.4 8.0 36.3 18.0 Propylene 0.0 10.3 7.9 9.617.0 17.10 Iso-butane 5.1 17.1 14.4 14.4 11.7 17.50 n-butane 3.7 15.612.9 12.9 8.4 13.2 Tr2-butene 1.2 1.3 1.3 1.1 0.70 0.10 1-butene 89.738.4 50.2 32.0 23.8 33.90 Iso-butene 0.10 0.0 0.10 0.0 0.0 0.0Cis2-butene 0.10 0.2 0.20 0.20 0.10 0.0 n-Pentene 0.0 9.7 7.5 7.6 1.90.40 3-hexene 0.0 1.7 1.3 13.8 0.10 0.0 97.7 i-C₆ 0.0 0.0 0.0 0.04 0.00.0 0.32 C7 & C8 0.0 0.0 0.0 0.25 0.0 0.0 2.0 Total, wt % 100 100 100100 100 100 100 Flow, Kg/h 4817 16,216 21,033 21,033 2,120 2696

TABLE 22 Material balance for the hexene isomerization section (5333.3hours operation) C₅ C₆ Component, C₅ Isomerization Isomerization HS11-hexene C₆ C₇₊ wt % Recycle Feed Prod Prod Prod Purge Purge 1-hexene2.4 2.2 8.4 92.7 98.5 2.4 0.3 Tr2-hexene 45.9 49.5 43.0 2.8 0.4 45.922.4 Tr3-hexene 21.4 20.0 20.1 2.0 0.3 21.4 9.5 Cis2-hexene 23.0 21.521.6 0.0 0.0 23.0 16.1 Cis3-hexene 7.1 6.6 6.6 2.2 0.5 7.1 2.8 i-C₆ 0.070.08 0.08 0.3 0.33 0.07 0.1 C₇ & C₈ 0.11 0.20 0.20 0.0 0.0 0.11 48.8Total, wt % 100 100 100 100 100 100 100 Flow, Kg/h 14,956 15,966 15,9661119 937 42 31

The comparison of the overall material balance of a continuous processwith the process of Example 2 indicated that the major streams scaledown linearly for the campaign case. Since the depropylenizer tower waseliminated, the C2/C3 to cracker in the campaign process was equivalentto the C2/C3 to cracker and depropylenizer side-draw purge streamscombined from the continuous case. This stream was slightly higher,requiring about 2% more raffinate II feed.

The energy balance for the campaign case study, producing 5 KTA ofpolymer grade 1-hexene is given below. The energy balance for buteneisomerization section, autometathesis section and hexane isomerizationsections are given.

TABLE 23 Energy balance summary for improved batch process C₄ Auto- C₆Isom metathesis Isom (2000 (2000 (5333 hours) hours) hours) Total Feedvaporizer (LPS), KW 1703 2364 1555 Feed Heater (fuel), KW 114 344 297Main Tower Condenser duty 13,950 9748 (CW), KW Main Tower Re-boiler duty11,480 7811 (LPS), KW Depentenizer Condenser duty 2057 5149 (CW), KWDepentenizer Re-boiler duty 1474 5147 (LPS), KW Depentenizer distillatecooler 1743 (CW) Pump, power, KW 98 50 90 Total Utility FUEL, KW 114 344297 CW, KW 13950 3800 14897 LPS, KW 13183 2364 14513 HPS, KW 1474 POWER,KW 98 50 90 Total Utility (8000 Hour Basis) FUEL, KW 313 CW, KW 14,369LPS, KW 13,562 HPS, KW 379 POWER, KW 97 FUEL, MKCAL/H 0.217 CW, MKCAL/H12.32 LPS, MKCAL/H 11.62 HPS, MKCAL/H 0.32 POWER, MKCAL/H 0.083

The energy consumption in the campaign process of Example 2 is increasedby the elimination of certain energy integration. In the improvedcampaign operation of this example, the energy efficient two-tower C₆separation system of Example 1 was replaced by a single tower separationto reduce the capital cost. Comparison of results in Tables 12 and 23shows that the cooling water and low pressure steam usage nearly doubledin the Example 2 case. This is due to the elimination of two-towerseparation system. The −5 deg C refrigeration is eliminated in theimproved batch process with a positive impact on operating cost. Example2 shows that the capital cost savings realized by these improvements inthe batch process as compared to Example 1 more than offset theincreased utility cost for the small-scale plant.

The utility summary before exchange for the 50 KTA continuous case withhigh selectivity catalyst is given above in Table 13. The 5 KTA improvedbatch utility consumption is very similar to this result on a linerprorated basis. The fuel usage decreased slightly as −5 deg C asrefrigeration is eliminated. The cooling water and LPS usage is higher.The capital cost savings realized by the equipment reduction in Example2 as compared to Example 1 more than offset the increased utility costfor the 5 KTA plant.

Example 3 Campaign Process Employing Heat Pump

Simulations were conducted in the steady-state process simulator HYSYSusing the PRSV property package. The analysis was carried out for a heatpump on the C₄/C₆ fractionator only, with conventional utilities used onthe depentanizer/C₆ fractionator.

In this example, the process that was used corresponded to that shown inFIG. 3 along with the heat pump shown in FIG. 4. The C₄ overhead fromfractionator 214 on FIG. 4 (corresponding to 114 on FIG. 3) contained 90mol % butene-1. The fractionator in the isomerization section contained100 theoretical stages and was operated at a reflux ratio of 29.9. Theoverhead pressure was 570 kPa, and the temperature profile for thefractionator operating in C₄ service is shown in FIG. 5. The ratio ofrecycle to fresh feed in the fractionator in the isomerization sectionwas 2.4 to 1. The isomerization reactor was operated at 343 Deg. C and2948 kPa with 21% conversion of butene-2 to butene-1. Autometathesistook place at 315 Deg. C and 1950 kpA. About 30% conversion of hexene-3was obtained. The depentenizer was operated at 1200 kPa with 30theoretical stages and a reflux ration of 1.0. The temperature profileof the depentenizer is shown in FIG. 6. The bottoms stream from thedepentenizer contained 98 mol % hexene-3. In the isomerization reactorabout 8.9% conversion to hexene-1 occurred. It is noted that condenser224 on FIG. 4 is condenser 117 on FIG. 3 and reboiler 206 on FIG. 4 isreboiler 123 on FIG. 3.

Using the 100 stage fractionator for C₆ processing, the fractionationreflux ratio for C₆ was 85.4 and the overhead pressure was 60 kPa. Theoverhead stream was separated to 92 mol % hexene-1. This stream was sentto the fractionator that previously had been run as a depentanizer. Theoverhead product obtained from the second fractionator, now in C₆service, was 8.5 mol % hexene-1. The fractionator operated at a refluxratio of 28.1 and an overhead pressure of 50 kPa. The bottoms stream wasrecycled to the isomerization reactor. Temperature profiles for thefirst and second fractionators when operated in C6 service are shown inFIGS. 6 and 7, respectively.

It was assumed that the compressor and expander were operated atconstant discharge pressures for both C₄ and C₆ services. To achievethis, the pressure of the C₆ fractionator was adjusted downwardly tobring its temperature profile within the range of that of the C₄fractionator. At the selected pressures of 570 kPa and 60 kPa for C₄ andC₆, respectively, the condenser and reboiler temperatures are given inTable 24.

TABLE 24 Condenser and Reboiler Temperatures C₄ Fractionator C₆Fractionator Equipment 570 kPa 60 kPa Condenser Temperature (C.) 47.7646.77 Reboiler Temperature (C.) 57.91 72.12

The limiting temperatures in Table 24 are the highest reboilertemperature and lowest condenser temperature. Therefore, the C₆fractionator determined both the compressor and expander dischargepressures. Using 3° C. outlet temperature approaches with the outlettemperature approach being defined as the difference between the processfluid outlet temperature and the heat pump fluid outlet temperature, thecompressor was required to raise the fluid boiling point to a minimum of75.12 C, and the expander was required to lower the boiling point to aminimum of 43.77 C. For n-butane as a heat pump fluid, the requiredpressures for these boiling point temperatures, shown in Table 25, were916.3 kPa and 420.9 kPa, respectively. Because the C₆ fractionator islimiting on both ends, the condenser and reboiler in C₄ service havetemperature approaches greater than 3° C., with the most notableapproach being 17.21° C. in the reboiler.

TABLE 25 Compressor and Expander Discharge Compressor ExpanderTemperature (C.) 75.12 43.77 Pressure (kPa) 916.3 420.9

With the compressor and expander pressures fixed, the outlet temperatureof heat exchanger 218 was set to the lowest achievable by cooling water,38° C. With the boiling temperature fixed, subcooling to 38° C. providedadditional cooling capacity, thus minimizing the fluid circulation ratethrough the heat pump. The circulation rate was then determined by thelarger heat duty, which was the condenser. Setting the condenser outletvapor fraction to one allowed the program to calculate the circulationrate, and setting the reboiler outlet vapor fraction to zero calculatedthe heat removed, which was the absolute heat duty difference betweenthe reboiler and condenser, by heat exchanger 213.

The C₄ fractionator condenser and reboiler duties were greater thanthose of the C₆ fractionator. Constant compressor and expander dischargepressures (and thus boiling points) were maintained with varyingfractionator duties by adjusting the heat pump circulation ratedepending on the service. Thus, the minimum circulation rate to fullycondense and vaporize the fluid throughout the cycle was employed. Theheat pump circulation rate was 2,393 kgmol/h for the C₄ fractionator and1,615 kgmol/h for the C₆ fractionator. Table 26 shows the circulationrates and heat duties for both towers.

TABLE 26 Heat Pump Data C₄ Fractionator C₆ Fractionator Circulation Rate(kgmol/h) 2,393 1,615 Condenser Duty (MW) 13.95 9.75 Reboiler Duty (MW)11.48 7.81 E-1 Duty (MW) 0.12 0.54 E-2 Duty (MW) 3.81 2.57

Analogous to the fractionator sizing for batch operation it is desirableto have C₄ and C₆ service require equal heat exchanger surface area inboth the condenser and reboiler. The difference in heat duties isconvenient, as the larger heat duties of the C₄ fractionator are alsoassociated with larger LMTD due to the temperature constraints imposedby the C₆ operation. The required heat exchanger surface area can beapproximated by calculating UA, given the heat duties and temperatureapproaches, for each exchanger. The calculation results are shown inTable 27.

${UA} = \frac{Q}{LMTD}$

TABLE 27 Heat Exchanger Sizing Condenser Reboiler C₄ C₆ C₆ FractionatorFractionator Depentanizer Fractionator Heat Duty (MW) 13.95 9.75 11.487.81 LMTD (C.) 6.5 5.4 17.2 5.8 UA (MW/C.) 2.14 1.74 0.67 1.26

It is noted that while neither the condensers nor the reboilers are anexact match, the difference in UA between C₄ and C₆ operation wascompensated for by using different numbers of heat exchanger shells inseries. For the cooling water exchangers in the heat pump loop, heatexchangers 213 and 218, the difference in duties was balanced by varyingthe cooling water flow rates.

In this example the minimum heat pump energy consumption was achieved inthe following ways:

-   -   With the constraint that the compressor and expander operate at        constant discharge pressures during C₄ and C₆ service, the        minimum compressor work was obtained by choosing the lowest        possible compressor discharge pressure and highest possible        expander discharge pressure. In this way, the least amount of        compressor work is undone by the expander. To set these limits,        the boiling temperatures of n-butane were chosen to be within        the minimum temperature approach (3° C.) of the condenser and        reboiler.    -   Cooling water was used in heat exchanger 218 to subcool to        38° C. before this stream entered the condenser. This maximized        the low-cost additional cooling capacity, thus reducing the        required circulation rate.    -   The minimum circulation rate was ensured by utilizing only the        latent heat of the n-butane in both portions of the heat pump        cycle. Heating or cooling into the region of sensible heat is        less efficient on a per mass basis, thus requiring a higher        circulation rate. In addition, only enough heat duty was removed        in heat exchanger 213 (tantamount to undoing compressor work) to        compensate for the difference in condenser and reboiler duties.        The circulation rate can be lowered in C₆ service to adjust for        the lower exchanger duties.

Example 4 Energy Consumption Analysis With and Without Heat Pump

In this example, energy consumption was simulated for the cases thatused and did not use a heat pump. When no heat pump is used, both theC₄/C₆ fractionator and depentanizer/C₆ fractionator have temperatureprofiles that permit the use of cooling water and steam in the condenserand reboiler, respectively. In the heat pump case it was assumed thatelectrical energy was required to drive the heat pump compressor.Alternately, high pressure steam can be used. The choice of compressorutility is dependent on many factors, such as cost, plant location, andavailability, and thus should be considered on a case-by-case basis. Thetower condenser and reboiler required no additional energy input.Cooling water can be used in the exchangers 113 and 118. Because theheat pump was applied only to the C₄/C₆ fractionator, the utilities ofthe depentanizer/C₆ fractionator number 2 were unchanged from theconventional case.

The analysis is summarized in Table 28. Energy costs were calculated andcompared for one year of batch operation with the process operating inC₄ service for 2,000 hours and in C₆ service for 5,333 hours.

TABLE 28 Utility Summary C₄ Service C₆ Service (2000 h) (5333 h)Conventional Case (Duties in MW) Equipment Utility Type C₄ FractionatorC₆ Fractionator Condenser Cooling Water 13.95 9.75 Reboiler Steam 11.487.81 Equipment Utility Type Depentanizer C₆ Fractionator #2 CondenserCooling Water 2.03 2.65 Reboiler Steam 1.47 2.64 Heat Pump Case (Dutiesin MW) Equipment Utility Type C₄ Fractionator C₆ Fractionator CompressorElectrical 1.60 1.08 Heat Exc. 113 Cooling Water 0.12 0.54 Heat Exc. 118Cooling Water 3.81 2.57 Second Equipment Utility Type Depentanizer C₆Fractionator Condenser Cooling Water 2.03 2.65 Reboiler Steam 1.47 2.64

From Table 28, the condensers of the C₄ fractionator and depentanizerused 13.95 and 2.03 MW of cooling water duty, while the two C₆fractionators used 9.42 and 2.65 MW, respectively. Cooling tower duty isvalued at $0.50/MBtu. For one year of operation:

${C\; 4\mspace{20mu}{Cost}} = {{\left( {13.95 + 2.03} \right){{MW} \cdot \frac{1\mspace{20mu}{Btu}}{1055\mspace{14mu} J} \cdot \frac{{\$ 0}{.50}}{MBtu} \cdot 2000}\mspace{20mu}{h \cdot \frac{3600\mspace{14mu} s}{h}}} = {{\$ 54}\text{,}500}}$${C\; 6\mspace{14mu}{Cost}} = {{\left( {9.75 + 2.65} \right){{MW} \cdot \frac{1\mspace{14mu}{Btu}}{1055{\mspace{14mu}\;}J} \cdot \frac{{\$ 0}{.50}}{MBtu} \cdot 5333}\mspace{14mu}{h \cdot \frac{3600\mspace{14mu} s}{h}}} = {{\$ 109}\text{,}800}}$

The reboilers of all but the depentanizer operated at low enoughtemperature to use low pressure (50 psig) steam, valued at $2.80/metricton. From saturated steam tables, 50 psig steam has a latent heat ofvaporization of 2121.6 kJ/kg.

${C\; 4{{Frac} \cdot {Cost}}} = {{11.48\mspace{14mu}{{MW} \cdot \frac{1000\mspace{14mu}{kW}}{MW} \cdot \frac{{\$ 2}{.80}}{1000\mspace{14mu}{kg}} \cdot \frac{1\mspace{14mu}{kg}}{2121.6\mspace{14mu}{kJ}} \cdot 2000}\mspace{14mu}{h \cdot \frac{3600\mspace{14mu} s}{h}}} = {{\$ 109}\text{,}100}}$C 6Frac ⋅ Cost = (7.81 + 2.64)$\mspace{65mu}{{{{MW} \cdot \frac{1000\mspace{14mu}{kW}}{MW} \cdot \frac{{\$ 2}{.80}}{1000\mspace{14mu}{kg}} \cdot \frac{1\mspace{14mu}{kg}}{2121.6\mspace{14mu}{kJ}} \cdot 5333}\mspace{14mu}{h \cdot \frac{3600\mspace{14mu} s}{h}}} = {{\$ 251}\text{,}600}}$

The depentanizer reboiler operated at 178.8° C., thus requiring mediumpressure (150 psig) steam, which has a saturation temperature of 185.6°C. Medium pressure steam was valued at $4.70/metric ton⁴ and has alatent heat of vaporization of 1994.9 kJ/kg.

${{DeC}\; 5\;{Cost}} = {{1.47\mspace{14mu}{{MW} \cdot \frac{1000\mspace{14mu}{kW}}{MW} \cdot \frac{{\$ 4}{.70}}{1000\mspace{14mu}{kg}} \cdot \frac{1\mspace{14mu}{kg}}{1994.9\mspace{14mu}{kJ}} \cdot 2000}\mspace{14mu}{h \cdot \frac{3600\mspace{14mu} s}{h}}} = {{\$ 25}\text{,}000}}$

Total annual utility cost when no heat pump was included was $550,000.For a batch system without the heat pump producing 5 KTA of hexene-1product, the utility cost is $0.110 per kilogram of hexene-1.

When a heat pump is included in the simulation, the heat pump compressorrequired 1.60 MW when operated on the C₄ fractionator, and 1.08 MW onthe C₆ fractionator. The power values were calculated using an adiabaticefficiency of 75%. Electrical energy was valued at $0.02 per kW-h⁴.

${C\; 4\;{Cost}} = {{1.60\mspace{14mu}{{MW} \cdot \frac{{\$ 0}{.02}}{kWh} \cdot \frac{1000\mspace{14mu}{kW}}{MW} \cdot 2000}\mspace{14mu} h} = {{\$ 64}\text{,}100}}$${C\; 6\;{Cost}} = {{1.08\mspace{14mu}{{MW} \cdot \frac{{\$ 0}{.02}}{kWh} \cdot \frac{1000\mspace{14mu}{kW}}{MW} \cdot 5333}\mspace{20mu} h}\mspace{11mu} = {{\$ 115}\text{,}300}}$

The cooling water required is that for E-1 and E-2. We take the samevalue as in the conventional case.

${C\; 4\;{Cost}} = {{\left( {0.12 + 3.81} \right){{MW} \cdot \frac{1\mspace{14mu}{Btu}}{1055\mspace{14mu} J} \cdot \frac{{\$ 0}{.50}}{MBtu} \cdot 2000}\mspace{14mu}{h \cdot \frac{3600\mspace{14mu} s}{h}}} = {{\$ 13}\text{,}400}}$${C\; 6\;{Cost}} = {{\left( {0.54 + 2.57} \right){{MW} \cdot \frac{1\mspace{14mu}{Btu}}{1055\mspace{14mu} J} \cdot \frac{{\$ 0}{.50}}{MBtu} \cdot 5333}\mspace{14mu}{h \cdot \frac{3600\mspace{14mu} s}{h}}} = {{\$ 28}\text{,}300}}$

The utility cost for the depentanizer/C₆ fractionator, which remainedthe same as the conventional case, was $122,900. When included with theC₄/C₆ fractionator, the total utility cost for the heat pump case was$344,100 per year, or $0.069 per kilogram of hexene-1 product.

The cost calculations are summarized below in Table 29. It is noted thatthe heat pump case generated savings of 61% of the conventional energyconsumption, and 37% of the energy cost. The savings are $0.041 perkilogram of hexene-1 product.

TABLE 29 Utility & Cost Summary C4 Service (2000 h) C6 Service (5333 h)No Heat Pump C4 Fractionator C6 Fractionator Equipment Utility Type Duty(MW) Annual Cost Duty (MW) Annual Cost Condenser Cooling Water 13.95$47,611 9.75 $85,707 Reboiler Steam 11.48 $109,091  7.81 $184,684  Total25.43 $156,703  17.56 $270,391  Depentanizer C6 Fractionator #2Equipment Utility Type Duty (MW) Annual Cost Duty (MW) Annual CostCondenser Cooling Water 2.03  $6,932 2.65 $24,064 Reboiler Steam 1.47$25,002 2.64 $66,941 Total 3.50 $31,934 5.29 $91,005 Average Heat Duty(MW) 25.30 Total Cost $550,033 ($0.110/kg hexene-1) Heat Pump C4Fractionator C6 Fractionator Equipment Utility Type Duty (MW) AnnualCost Duty (MW) Annual Cost Compressor Electrical 1.60 $64,067 1.08$115,311  E-1 Cooling Water 0.12   $417 0.54  $4,939 E-2 Cooling Water3.81 $13,005 2.57 $23,399 Total 5.53 $77,488 4.19 $143,650  DepentanizerC6 Fractionator #2 Equipment Utility Type Duty (MW) Annual Cost Duty(MW) Annual Cost Condenser Cooling Water 2.03  $6,932 2.65 $24,064Reboiler Steam 1.47 $25,002 2.64 $66,941 Total 3.50 $31,934 5.29 $91,005Average Heat Duty (MW) 9.69 Total Cost $344,077 ($0.089/kg hexene-1)Heat Duty Savings 61.7% Cost Savings 37.4% ($0.041/kg hexene-1)

For one year of campaign operation producing 5 KTA of hexene-1, theenergy consumption and costs of the two cases, with and without the heatpump, are compared in Table 30. Inclusion of the heat pump saves 61% ofthe total duty and 37% of the utility cost of the conventional case.

TABLE 30 Utility Summary Conventional Case Heat Pump Case Avg DutyAnnual Avg Duty Annual Utility Type (MW) Cost (MW) Cost Cooling Water13.1 $164,314 5.8 $72,756 Steam 10.8 $385,719 2.3 $91,943 Compressor 0.0$0 1.2 $179,378 Total 23.9 $550,033 9.4 $344,077 Savings — — 60.8% 37.4%

In summary, the campaign process is different from the prior-knowncontinuous process in the following ways:

1. The C₄ and C₆ fractionator/isomerization reactor systems are designedas a single unit. The system is operated for one period of time periodof time as a C₄ isomerization/fractionation system and during anotherperiod of time as a C₆ isomerization/fractionation system.

-   -   2. With campaign processing, the combined        fractionator/isomerization reactor system is shared. An        intermediate storage tank is required to allow for this type of        operation. In the batch system, the fractionator/isomerization        reactor operates first in C4 service to produce butene-1. This        butene-1 can go to product butene-1, go further to        autometathesis, or both. The autometathesis effluent produces        lights for recovery, a recycle stream of C4/C5 olefins, and        hexene-3 to fill a storage tank. The fractionator/isomerization        reactor is then converted to C6 service to produce hexene-1    -   3. The third modification, in certain embodiments, is the        inclusion of the closed loop heat pump with the working heat        transfer fluid composition set to match both C4 and C6        operations. Using the same closed loop system for two different        carbon number systems is unique. By operating in the campaign        mode, the utility saving features of the cross exchange between        the butene and hexene superfractionators is unavailable. This        will result in higher utilities per unit hexene-1 product        compared to the improved continuous process. To offset the added        utility costs, a circulating heat transfer stream is alternately        compressed and expanded to adjust its boiling point within the        temperature range of the reboiler and condenser. The closed loop        heat pump is used in these exchangers in place of conventional        utilities. Flow rates and operating times are manipulated such        that the design of the fractionator/isomerization reactor        combination and the heat pump is amenable to both C4 and C6        service for use with the batch system.    -   4. In order to further reduce capital costs, the depentanizer is        used as a topping column for hexene-1 purification. The        fractionation duty required to produce high purity butene-1 is        less than that required to produce hexene-1 (from their        respective isomers). Thus a tower designed for both services has        to be “oversized” for butene service to accommodate the hexene        service. However, during the hexene-1 operation, the        depentanizer tower is not in service in certain embodiments.        Thus the fractionation capability of this tower can be used to        provide the additional fractionation capacity for the hexene-1        purification thus allowing the main fractionator to be sized for        butene service, resulting in additional capital savings.

The campaign processes described herein provide benefits over theconventional continuous process in three ways. First, the sharedequipment of the campaign process reduces the total capital cost, asopposed to requiring dedicated equipment for C₄ and C₆ service. Whilethe addition of the heat pump adds cost to the campaign process itself,its use results in a reduction in utility costs. Second, the campaignnature of the process and the flexibility to vary the operating timesfor each use of the isomerization technology allows for variation inproduction of butene-1 and/or hexene-1 depending upon changing marketconditions. Third, the heat pump addresses the increase in energyconsumption incurred by recycle of the isomerization effluent, affectinga comparable utility cost to that of a lower-yield process which forgoesthe isomerization system.

It will be appreciated that various of the above-disclosed and otherfeatures and functions, or alternatives thereof, may be desirablycombined into many other different systems or applications. Furthermore,it is noted that presently unforeseen or unanticipated alternatives,modifications, variations or improvements therein may be subsequentlymade by those skilled in the art which are also intended to beencompassed by the following claims.

What is claimed is:
 1. A system for producing an alpha olefin,comprising: a first isomerization reactor configured to isomerize afirst batch of an olefin having a first carbon number to form a firstisomerization reactor effluent and subsequently process a second batchof an olefin having a second carbon number to form a secondisomerization reactor effluent, a metathesis reactor positioneddownstream from the first isomerization reactor, the metathesis reactorbeing configured to disproportionate the first isomerization reactoreffluent to form a metathesis reaction product, a first fractionatorpositioned downstream from the isomerization reactor and beingconfigured to separately fractionate the first and second isomerizationreactor effluents, a second fractionator positioned downstream from themetathesis reactor to remove light hydrocarbons from the metathesisreaction product, a storage tank disposed downstream from the first orsecond fractionator, and a storage tank outlet line connecting thestorage tank to an inlet of the first isomerization reactor and/or tothe inlet of the metathesis reactor.
 2. The system of claim 1, whereinthe first fractionator includes two separate fractionation columns. 3.The system of claim 1, wherein the second fractionator is configured toalso provide additional fractionation capacity for fractionating thesecond isomerization reactor effluent.
 4. The system of claim 1, whereinthe first fractionator and/or the second fractionator are part of afractionation sub-system including a condenser and a reboiler, and thecondenser and reboiler form a heat pump.
 5. The system of claim 4,wherein the heat pump further includes a vaporizer associated with thefractionator.
 6. A system for sequentially producing butene-1 andhexene-1, the system comprising: an autometathesis reaction system forconverting 1-butene in a 1-butene fraction into a metathesis effluentcomprising C2-C7 hydrocarbons including ethylene and 3-hexene; anisomerization reactor loop configured to be used for both: isomerizing aportion of 2-butene in a 2-butene fraction to form a C4 isomerizationproduct comprising 1-butene, and isomerizing a portion of the 3-hexenein a 3-hexene fraction to form a C6 isomerization product comprising1-hexene; a fractionation system configured to be used as both: a buteneseparation system for separating a C4 feed and the C4 isomerizationproduct into the 1-butene fraction and the 2-butene fraction, and ahexene separation system for separating the C6 isomerization productinto a first fraction comprising 1-hexene and a second fractioncomprising 3-hexene.
 7. The system of claim 6, further comprising atleast one of a 1-butene fraction storage tank and a 3-hexene fractionstorage tank.
 8. The system of claim 6, further comprising a metathesiseffluent fractionation system for separating the metathesis effluentinto one or more hydrocarbon fractions including the 3-hexene fraction.9. The system of claim 8, wherein the metathesis effluent fractionationsystem comprises: a depentanizer for separating the metathesis effluentinto an overheads fraction comprising C5 and lighter hydrocarbons and abottoms fraction comprising C6 and heavier hydrocarbons; adepropylenizer for separating the overheads fraction into a C2/C3overheads fraction, a C4/C5 bottoms fraction, and, optionally, a highpurity 1-butene side draw fraction.
 10. The system of claim 9, furthercomprising a steam cracker separation system configured to receive atleast a portion of the C2/C3 overheads fraction.
 11. The system of claim9, further comprising a recycle line for feeding at least a portion ofthe C4/C5 bottoms fraction to the autometathesis reaction system. 12.The system of claim 6, wherein the fractionation system comprises afirst tower and a second tower, the first tower configured to operate ata higher pressure than the second tower.
 13. The system of claim 12,further comprising a C4 raffinate feed line for introducing the C4raffinate feed to the second tower, a first isomerization product feedline for introducing a first portion of the C4 or C6 isomerizationproduct to the first tower, and a second isomerization product feed linefor introducing a second portion of the C4 or C6 isomerization productto the second tower.
 14. The system of claim 13, further comprising aheat exchange system for supplying heat removed in an overhead condenserof the first tower to a reboiler of the second tower.
 15. The system ofclaim 6, further comprising a purge line for removing heavies from theC6 isomerization product and the C4 isomerization product processed inthe first tower.
 16. A system for sequentially producing 1-hexene and1-butene, the system comprising: an isomerization reactor loopconfigurable for C4 and C6 service, where: in C4 service, theisomerization loop is configured to (i) receive at least one of a C4raffinate fraction comprising 2-butene, and a fractionation productcomprising 2-butene and (ii) convert a portion of the 2-buene to1-butene producing a C4 isomerization effluent; and in C6 service, theisomerization loop is configured to (i) receive a 3-hexene fraction anda recycle C6 fraction comprising 2-hexene and 3-hexene, and (ii) converta portion of the 3-hexene and 2-hexene to 1-hexene producing a C6isomerization effluent; a first fractionation system configurable for C4and C6 service, where, in C4 service, the fractionation systemfractionates the C4 isomerization effluent into a 1-butene fraction andthe fractionation product comprising 2-butene; in C6 service, thefractionation system fractionates the C6 isomerization effluent into anoverhead fraction of increased 1-hexene purity, relative to the C6isomerization effluent, and a first portion of the recycle C6 fraction;an autometathesis reaction system for converting 1-butene in the1-butene fraction into a metathesis effluent comprising C2-C7hydrocarbons including ethylene and 3-hexene; a second fractionationsystem configurable for C4 and C6 service, where: in C4 service, thefractionation system fractionates the metathesis effluent to produce aC2/C3 fraction, a C4/C5 recycle fraction, and the 3-hexene fraction; inC6 service, the fractionation system fractionates the hydrocarbonfraction into a 1-hexene product fraction and a second portion of therecycle C6 fraction.
 17. The system of claim 16, wherein the secondfractionation system comprises a flash drum and a distillation column,the distillation column configured to operate as a depentanizer when thesecond fractionation system is configured for C4 service, and the flashdrum being used only when the second fractionation system is configuredfor C4 service for separating the depentanizer overheads into the C2/C3fraction and the C4/C5 recycle fraction.
 18. The system of claim 16,further comprising a 3-hexene fraction storage tank.
 19. The system ofclaim 16, further comprising a steam cracker separation systemconfigured to receive at least a portion of the C2/C3 fraction.
 20. Thesystem of claim 16, further comprising a recycle line for feeding atleast a portion of the C4/C5 recycle fraction to the autometathesisreaction system.